Multiple Reactor Chemical Production System

ABSTRACT

The present invention is a multiple reaction set for the production of chemicals by equilibrium limited reactions utilizing plate-type or extended surface heat exchangers. The heat exchangers cool the reaction products to condense methanol within the reaction products for separation, and to warm incoming feed reactants prior to entrance of the reactants into a reactor utilized for methanol production. The various reactors, heat exchangers, and separators can be formed as separated zones within enclosed vessels, thereby eliminating the need for separately constructed reactors, heat exchangers, and separators. Multi-stream plate-type of extended surface heat exchangers can be utilized to allow efficient cooling and methanol separation. The multiple reaction set can also be used for the recovery of methanol from a waste or purge gas stream utilizing multiple reactors, multiple plate-type or extended surface heat exchangers and multiple separators as a substitute for or in conjunction with a conventional methanol synthesis loop.

CROSS-REFERENCE TO RELATED APPLICATIONS

This is a divisional application of U.S. patent application Ser. No.12/067,608, filed on Aug. 6, 2008, which claimed priority under 35U.S.C. §119(e) to U.S. provisional patent application Ser. No.60/720,330, filed on Sep. 23, 2005 the entirety of which are expresslyincorporated by reference herein.

BACKGROUND OF THE INVENTION

It is rare in a chemical process that the process uses stoichiometricamounts of reactants with essentially complete conversion in a simplereactor. Hence, where the reagents constitute a significant cost of theprocess, the unreacted material is often recycled to the reactor,usually after some physical separation of the desired product from theun-reacted material. Sometimes this separation can be achievedinternally within the reactor, for example, where the reactants aregaseous, the product is in liquid form at reaction conditions and iswithdrawn continuously, and a stirred tank reactor with gas inductionimpellers is utilized. In this particular situation, the physicalseparation and the recycle occur within the reactor vessel.

Alternatively, the separation and recycle of reactant can take placeexternal to the reactor. One example of this configuration would be aplug flow gas phase reactor where the product can be condensed from thegas phase by cooling. The unreacted gases may then be re-compressed, andat least partially returned to the inlet of the reactor, perhaps afterother conditioning, such as purification or chemical separation.

There are several reasons why the amounts of reactants utilized in thereactor to form the desired end product are rarely stoichiometric. Itcould be that vapor pressure limitations require a non-stoichiometricreaction. For example, in high-pressure, gas-phase hydrogenations of ahigh boiling organic the hydrogen will be present in large excess eventhough a high single pass conversion of the reactant is theoreticallypossible. An alternate reason would be that the reaction is equilibriumlimited. For example in acid catalysed esterifications the alcohol isfrequently in excess to achieve high conversion of the acid.

While it may be possible to achieve high conversion of the reactants inan equilibrium limited reaction to the desired end product economicallyutilizing a large excess of reactants, an alternate possibility is theremoval of one of the reaction products. For example, gas strippingcould remove the water from an esterification reaction to continuallymove the conditions in the reactor out of or away from equilibrium andthereby drive the reaction toward full conversion.

However, where a product cannot be removed in-situ to drive theequilibrium toward full conversion, then a high overall conversion islikely to be achieved only by separation of the product from thereactant mixture and subsequent recycle of the un-reacted material tothe reactor.

Additionally, even where an equilibrium reaction has certain conditionsor aspects that are favorable for high conversion to the desired endproduct, the kinetics of the reaction may suggest that a higher overallproduction rate or better process economics can be achieved by running areactor at conditions favoring a relatively low conversion of thereactants and then recycling the un-reacted material after physicalseparation of the product. Some exemplary reactions or processes of thistype where conversion of the reactants to the desired product is onlypartial, and in which significant quantities of un-reacted materialremain that can be recycled to the reactor after physical separation ofthe product include the reaction of synthesis gas to methanol, di-methylether, mixtures thereof, Fischer-Tropsch waxes, and ammonia.

Using the equilibrium limited reaction for the production of methanol asan example, because methanol is one of the largest by volume chemicalsproduced in the world today, the conversion to methanol is typicallycarried out in a two-step process. In a first step, methane is reformedwith water or partially oxidized with oxygen to produce carbon monoxideand hydrogen, with some carbon dioxide and residual methane, (i.e.,synthesis gas or syn-gas). In a second step, the syn-gas is convertedinto methanol.

The second step of converting the syn-gas into methanol is a well-knownprocess. It typically involves a catalytic process using a copper-basedcatalyst, such as a catalyst comprising a reduced zinc oxide/copperoxide mixture, among others. To provide the optimum production ofmethanol from this reaction, the reaction is typically carried out atpressures within the range of 40-100 bars and at temperatures in excessof 200 degrees C. and below 320 degrees C., with a temperature range ofbetween 220 and 280 degrees C. being most common. The production of thesyn-gas itself is typically carried out at pressures within the range of20-40 bars depending on the reformer technology that is utilized.

Due to the particular mechanism of the reaction for the production ofmethanol, the reaction does not go to completion, as the concentrationof produced methanol is limited by equilibrium. Specifically, the amountof methanol contained in the product gas exiting the reactor comprisesabout 6-8 mol % of the total gas, although it can be higher. Thismethanol is removed from this product gas stream by condensing itthrough cooling the product gas stream to below 110 degrees C., and mostcommonly below 60 degrees C. The cooled methanol can then be removedfrom the gas stream while the excess syn-gas is sent back to the reactorin order to further react the excess syn-gas. This enables additionalmethanol to be obtained from the syn-gas recycled back to the reactor incombination with an amount of fresh syn-gas that is also charged to thereactor.

In performing this recycle step, one well-known process involves the useof a recycle compressor which receives the excess syn-gas from theseparator and compresses it in order to overcome the pressure drop thatoccurs within the reactor and separator. This type of reactor iscommonly referred to as a recycle loop reactor and is schematicallyshown in FIG. 1. In this reactor, the concentration of methanol in thesyn-gas leaving the reactor is low enough that the volumetric flow rateof the excess syn-gas through the recycle compressor is typically two(2) to ten (10) times the volumetric flow rate of the fresh syn-gasbeing introduced into the reactor separately from the syn-gas charged tothe reactor from the recycle compressor. A purge gas stream ofapproximately 4 to 8 percent of the recycled syn-gas stream is alsoexpelled from the recycle loop, prior to re-compression to control theconcentration of inert material that builds up in the reactor as aresult of the recycle.

One significant drawback to the recycle loop reactor described above isthe cost of the recycle compressor. Often the recycle function isincorporated into a single drive train compressor that compresses thesyn-gas up to the pressure of the recycle loop and also provides forre-compression of the recycle gases. The compression train is anexpensive item of equipment and may be the single most expensivepurchased component in the construction of a methanol producingfacility. As stated previously, a high recycle flow is utilized toenable high overall conversion of syn-gas to be achieved. The recyclecompressor also becomes significantly cheaper per unit volume compressedas the scale of the plant is increased. Thus, the use of a process usinga recycle compressor is, as a practical matter, preferential for thosefacilities producing a relatively high daily output of methanol relativeto the current maximum single train production plant, such as thosefacilities producing in the neighborhood of 5,000 tons per day in 2005,where maximum efficiency is required and integration of the syn-gascompressor and recycle compressor can be achieved in order to make theuse of a recycle compressor economically viable.

As an alternative to a recycle loop reactor, it has been conceived thatthe facility could utilize a multiple reactor set or cascade process(FIG. 2) whereby the syn-gas is initially fed to a first reactor forreaction with the catalyst contained therein to produce methanol. Theproduct gas then enters a first separator or knock-out pot wherein themethanol produced in the first reactor is cooled into liquid form andseparated from the excess syn-gas. The first separator can operate toseparate the methanol for the syn-gas in any desired manner, such as bygravity or by applying centrifugal focus to the products. The remainingexcess syn-gas is then fed to a second reactor, which undergoes the samereaction, thereby producing additional methanol. The additional methanolis removed from the second reactor and directed to a second separator inthe same manner. The number of reactors and separators can be selectedto create a multiple reactor set that achieves the desired conversionpercentage of the syn-gas to methanol. For example, when using anoptimal syn-gas composition that has a conversion to methanol of 50percent in each reactor, the reactor set can be selected to include fourreactors and four separators, which theoretically results in theachievement of a 95 percent overall conversion of the syn-gas tomethanol after the fourth separator.

However, in more realistic situations, when the syn-gas composition isless optimal, such as when the stoichiometry of the syn-gas is away fromthe stoichiometry of the reaction desired, e.g., the ratio ([moleshydrogen]−[moles carbon dioxide])/([moles carbon monoxide]+[moles carbondioxide]) is between 2.5 and 3.0, then more than four reactor sets andperhaps as many as 10 reactor sets are required for high conversion(>95%). Additionally, when there are stoichiometric quantities ofreactant in the syn-gas, but high levels of inerts also present, then ahigh number of reactor sets will be required, such as where autothermalreforming with air is performed for the production of syn-gas resultingin dilution of the syn-gas with high levels of nitrogen.

There are several areas in a methanol production process where theability to employ a cascade of reactors would be considered beneficial.Principally the benefit from using the cascade comes from not requiringa recycle compressor. Furthermore, it should be borne in mind that in agrass roots methanol production plant the recycle compressor is oftenpart of the compressor associated with syn-gas compression. Thus,eliminating the syn-gas compressor along with the recycle compressor ina grass roots installation gives the maximum benefit.

While the use of a simple cascade of reactors for this particularpurpose is disclosed in certain prior art references, the referencesthat discuss the use of such a cascade of reactors focus exclusively onmethods by which the reformer operating pressure can be matched to themethanol synthesis pressure. For example, U.S. Pat. Nos. 5,177,114;5,245,110; 5,472,986; and 7,019,039 are each patents that discloseinventions in the field of autothermal reforming using air rather thanoxygen. However, while these patents very generally disclose the use ofcascade reactors in the methanol production process, they do not addressthe issues of how the cascade of reactors can be made cost effectively.Furthermore, each of U.S. Pat. Nos. 5,177,114; 5,245,110; and 5,472,986disclose a methanol production process where the recycle compressor canbe eliminated as a result of operating a reformer autothermally, andthen converting the syn-gas to a methoxy compound using three to fivereactor sets with product condensation between each stage. Recognizingconventional wisdom that a cascade process cannot achieve a high syn-gasto methanol conversion, carbon efficiencies for the methanol synthesissection of less than 80% are quoted, whereas in conventional plantefficiencies in excess of 95% are achievable.

Additionally, U.S. Pat. No. 6,255,357 discloses a methanol productionprocess that uses pressurization of the oxidant gas for fired heating ofthe steam reformer as a means of achieving a mechanically feasible highpressure steam reformer with an operating pressure sufficient to ensurea sufficient operating pressure throughout the process. The process alsoincludes a cascade of reactors downstream from the reformer in which thereformed syn-gas is converted to methanol. The pressurization of theincoming natural gas into the reformer avoids the requirement of asyn-gas compressor upstream of the cascade of reactors, as well asavoiding the need for a recycle compressor. However, as with theprevious references, the cascade of reactors is only very generallydisclosed without any discussion as to how the cascade can be madeeconomically.

Other situations regarding a methanol production process where it wouldbe considered advantageous to avoid use of a recycle compressor includethose where the compressor would be added to an existing methanolproduction facility in a retro-fit capacity, or as part of an additionto a planned methanol production facility construction for the purposeof removing any remaining methanol from the purge gas discharged fromthe facility. One system of this type that addresses the loss of thepotential and actual methanol present in the purge gas stream isdisclosed in U.S. Pat. No. 6,258,860, which is incorporated by referenceherein in its entirety. The process disclosed directs the purge gasstream produced by a methanol synthesis zone to another methanolsynthesis or production zone in order to both collect the methanolpresent in the purge gas stream as well as to further react theunreacted components of the purge gas stream to produce additionalmethanol.

However, the process disclosed in the '860 patent has certain drawbacksin that it utilizes a compressor to compress the combined purge gas andrecycled syn-gas stream prior to further reacting the combined stream.Because, as discussed previously, the recycle compressor is the highestcost item in a methanol production system, the use of additional recyclecompressors to recover methanol from a purge gas is highly undesirable,especially for systems producing a relatively low daily output ofmethanol relative to the current maximum single train production plant,such as those facilities producing in the neighborhood of 5,000 tons perday in 2005.

Another example where additional compressor capacity may be avoidedthrough the use of a set of cascade reactors could occur as part of are-vamp or de-bottle-necking of a methanol plant. If the re-vamp orde-bottlenecking entails increased syn-gas production, then the capacityof the methanol converter would be required to be increased. It may bepossible to increase the effectiveness of the reactor through betterpacking of catalyst or dividing the catalyst into multiple beds in asingle reactor. However, where the reactor already makes effective useof the catalyst, it may not be possible to economically increase theperformance of the reactor. Further, another limitation on the operationof the reactor in this situation is the pressure drop of the process gasacross the reactor. Increasing the re-circulation rate, increasing thecatalyst volume, increasing the feed flow or reducing the purge ratewill all increase the pressure drop of the process gas through thereactor. There will of course be a consequent limitation on the capacityof the re-circulation compressor as well to recompress the gas forreintroduction into the reactor.

One alternative to mitigate these issues would be to operate themethanol converter at reduced conversion conditions, but with a highergas feed rate, thereby allowing pressure drop limitations to be avoidedand then utilizing a separate cascade reactor system to convert theun-reacted gases to methanol without the requirement of an additionalcompressor or replacing the original reactor. This also has theadvantage of being a lower risk method of increased throughput, as theoriginal reactor performance is well known. There will of course be manyother circumstances under which a cascade system can be utilized, butall of these circumstances will be reliant on a cost-effective design ofcascade system.

One significant drawback with the multiple or cascade reactor set typesutilized in methanol production as described in the prior art resultsfrom the end use construction of each reactor, heat exchanger(s) andseparator forming the individual reactor set. Specifically, because aportion of the syn-gas is lost in each reactor set based on itsconversion to methanol, often each subsequent reactor set and separatoris constructed to be smaller than the immediately preceding ones toaccommodate the reduction in the flow rate of incoming syn-gas. Thismight initially be anticipated to be highly beneficial based upon thereduction in the amount of material necessary to construct eachsuccessive reactor set. However, each reactor set requires the samefunctionality, connections, cooling and access for catalyst replacement,which become more difficult and/or expensive to manufacture on aprogressively smaller scale. In addition, the cooling, gas-liquidseparation and re-heating of the methanol-bearing stream as it passesbetween the various reactor sets must be effected in an energy efficientand cost effective manner. Further, all of the reactor sets andseparators must be constructed to be operable at the elevated pressures(40-100 bars) that the reactions occurring for the conversion of thesyn-gas to methanol require.

One example of a system that attempts to address this shortcoming isdisclosed in U.S. Pat. No. 6,723,886 in a methanol production processusing reactive distillation. However, while there is removal of methanolbetween reactor beds by condensation within the reactor, thecondensation takes necessarily takes place at reaction temperature, andcondensation at elevated temperature limits the conversion of methanolto approximately 60%. However, even with the significant restrictionthis places on methanol production, this is in accord with the currentindustry view that condensation at reduced temperature is not viable.

Therefore, it is desirable to develop a multiple or cascade reactor setand a process for the production of products of equilibrium limitedreactions, e.g., methanol, using the multiple reactor set to obtain ahigh percentage conversion of feed syn-gas to methanol by condensing themethanol in the reactor effluent in an interstage feed/effluent heatexchanger. It is also desirable that the multiple reactor set beoperable without the need for a gas recycle compressor and preferablywithout the need for the construction of multiple individual reactors,heat exchangers, and separators. In other terms, the heat exchangerdesign should be suitable for efficient operation and integration intothe reactor sets, while also minimizing the number of necessaryequipment items.

With regard to the goal of minimizing the necessary number of equipmentitems in a reactor set, it is easier to understand the conventionalapproach to solving this problem of eliminating equipment items andreducing the cost of equipment items by reference to the specificproblems of a conventional methanol synthesis loop. Apart from therecycle compressor, a methanol synthesis loop contains six principleoperations: 1) pre-heat of the gas; 2) reaction of the gas to formmethanol; 3) removal of the heat of reaction as high grade heat; 4)cooling of the gas to methanol condensation temperatures; 5)condensation of the methanol using cooling water; and 6) vapor/liquidseparation. In a typical plant there may be two integrations of thesefunctions for the purposes of minimizing the necessary equipment itemswhich are removal of the heat of reaction is performed by steam raisingin a shell and tube reactor, and pre-heat of the gas by feed-effluentexchange. Thus, a typical synthesis loop will consist of at least sixequipment items: 1) a start up heater; 2) a feed/effluent heatexchanger; 3) a reactor; 4) a high grade heat recovery unit; 5) a watercooler; and 6) a gas-liquid separator.

Steam raising directly in the reactor does eliminate the requirement fora separate high-grade heat recovery unit. However, it also requires asteam drum with the reactor and so does not reduce the number ofequipment items.

With regard to the use of the feed/effluent heat exchanger, the highestenergy efficiency is achieved with a high effectiveness heat exchangerthat is able to maximize the cooling of the effluent stream. Increasingthe amount of high grade heat recovered reduces the temperaturedifference in the feed/effluent exchanger. Therefore, for maximum highgrade heat recovery a high effectiveness heat exchanger is required.However, shell and tube heat exchangers as used in prior art multiplereactor sets can only achieve high effectiveness through the coupling ofmultiple heat exchanger units, again increasing the number of equipmentitems required. The usefulness of high grade heat recovery is, in partdependent, on the temperature at which it is recovered. In particular,for a methanol process the high grade heat recovery from the methanolsynthesis section is utilized for steam raising for the reformer. Thisrequires that the stream from which heat is being recovered is above aminimum temperature, typically 200-250 deg C. However methanolcondensation temperatures are in the region of 60-100 deg C. Forefficient operation, therefore, heat exchangers are required that canoperate with a hot gas temperature span of approximately 150 deg C. Theheat of reaction is recovered by cooling the reactant stream by theequivalent of typically 50-100 deg C. of sensible heating. If the feedgas is introduced to the methanol reactor at a temperature below highgrade heat recovery temperature this represents a loss of energyefficiency in the system and increases the low grade coolingrequirement. Consequently the temperature difference in thefeed-effluent exchanger will be kept to less than 50 deg C. andtypically 20-30 deg C. Where high single pass conversions can beachieved in the reactor, such as with a balanced stoichiometry, highoperating pressure, efficient heat removal or a low overall conversionthe temperature constraints may be more relaxed. However, this oftenbrings greater reactor complexity or lower overall efficiency.

The performance measure of a heat exchanger can be described in terms oftemperature span and log mean temperature difference between stream. Thevalue (span divided my lmtd) is referred to as NTU count and as can beseen above it would be desired for an energy efficient methanol processthat the fee/effluent exchangers would operate with an NTU count above5, and more preferably above 7.

The problem concerning the number of equipment items is also notalleviated when a recycle loop is replaced with a cascade system. Withno recycle there is no requirement for a recycle compressor. However,for each contact with the catalyst there will be up to six additionalequipment items, as discussed previously. One option to reduce thenumber of equipment items is to eliminate some of the heat exchangers.For example, instead of recovering high grade heat from the reactorgases, the gases can be used to directly heat the incoming feed gases.The feed/effluent exchanger is then smaller as a result of an increaseddriving temperature, but the reaction heat is then lost to the coolingwater and a less efficient process is produced.

Therefore, to improve the economics and efficiency of the prior artmethanol cascade systems, it is necessary to solve the followingissues: 1) to minimize the number of equipment items; 2) to increase theeffectiveness of the feed/effluent exchangers; and 3) to integratemultiple functions into single equipment items.

SUMMARY OF THE INVENTION

One method by which the improvement of the economics and efficiency ofthe cascade reactor system can be achieved, and which is an integralpart of the apparatus and method of the present invention, is the use ofextended surface, or plate style heat exchangers in the cascade reactorsystem. In particular, plate-fin (brazed or diffusion bonded) or printedcircuit heat exchangers (PCHE) are able to achieve high effectivenesseliminating the requirement for multiple units for a single duty or apre-heater to reduce the required effectiveness. The constructionmethods of plate style heat exchangers also allow for multiple streamheat exchangers to be combined into a single unit. For example, thereactor outlet gases can pass through a single plate style heatexchanger where, in a first section, the gas is cooled with a hightemperature coolant such as pressurized water at 200-250 deg C. In asecond section of the heat exchanger, the reactor gases are cooled bythermal contact with the reactor inlet stream. Finally, in a thirdsection, the reactor gases are cooled with cooling water to condense thewater.

When a series or cascade of reactors is used, as discussed previously,the appropriate design for heat recovery may differ as the reactordiminishes in size. In the early reactors of the set the cost of heatrecovery is more economical as the energy recovered per unit isgreatest. With each successively smaller reactor and associated heatexchanger, the amount of energy that is available per unit diminishes asthe rate of production of methanol in the set is smaller and the costsof heat recovery can become prohibitive. So, the present invention canutilize as the first reactor a steam raising or gas tube cooled reactorwithin a recycle loop, as discussed previously. The subsequent reactorsets each utilize a high efficiency heat exchanger of the aforementionedtype integrating high grade heat recovery, feed effluent heat exchangeand methanol condensation utilizing cooling water, whereas the finalreactor set only utilizes a high efficiency heat exchanger forfeed/effluent heat exchange and water cooling.

For even smaller methanol production units, it may also be possible toincorporate the heat exchange functions, i.e., the high efficiency heatexchangers, for each reactor set into a single fabricated unit. This ismade possible by the use of plate style heat exchangers that areamenable to such construction. Additionally, separate from theintegration of the heat exchangers into a single unit, integration ofthe reactors themselves and also the vapor/liquid separators into thesame or separate units are also possible. At this smaller scale ofproduction, the present invention can have all of the reactors containedwithin a single item of equipment, and all of the heat exchangers in asingle unit, as well as all of the vapor/liquid separators. As a result,the cascade methanol process is effectively reduced to only threeprinciple equipment items.

The integration of these units allows for several different uses of theintegrated units. For example, a cascade reactor set unit of 3 or 4reactors could be connected to the purge gas stream from a methanolsynthesis loop to allow for the further reaction of the contents of thepurge gas stream through the unit to form additional methanol. Thiswould increase the overall conversion percentage of the loop withoutincreasing the recycle rate, as only the purge gas is directed throughthe cascade reactor set unit. Also, the addition of the cascade reactorunit would both increase the amount of methanol that can be made from afixed stream of methane, or, as part of a wider retro-fit to an existingproduction process, would boost the capacity of the methanol synthesissection without an increased gas rate through the recycle compressor.

Therefore, according to a first aspect of the present invention, animproved multiple or cascade reactor set type chemical production systemis provided in which the incoming reaction gas entering the firstreactor is an amount of excess reaction gas and/or a purge gas from aconventional single reactor, a cascade reactor set or a recycle loopreactor system, each of which are fed from a steam reformer or anautothermal reformer. The gas initially passes through a firsthigh-efficiency plate-type or extended surface heat exchanger, wherebythe incoming syn-gas comes into thermal contact with some or all thegaseous reaction products exiting the first reactor in order to cool thegaseous reaction products and preheat the incoming syn-gas. Some or allof the gases exiting the first reactor may be brought into thermalcontact with an additional stream for the recovery of high grade heatprior to entering the feed-effluent heat exchange section. The coolingof the reaction products causes the desired product to condense intoliquid form within the heat exchanger. Further cooling of the productbearing stream is then effected by thermal contact with an additionalcooling utility stream that is also introduced into the aforementionedheat exchanger, such an arrangement being known as a multi-stream heatexchanger. The cooled reaction products, including the condensed liquidproduct, can be removed directly from the heat exchanger or can thenflow into the first separator where the condensed liquid product isremoved in the separator while the excess reaction gas is directed intoa second reactor for additional production of the desired chemical.Prior to reaching the second reactor, the excess reaction gas passesthrough a second high-efficiency heat exchanger to be warmed by thegaseous reaction products exiting the second reactor, and consequentlycondenses the product contained in the product gases exiting the secondreactor. The high-efficiency plate-type or extended surface heatexchangers condense the desired product, for example methanol, producedin each of the reaction zones in a highly economical manner, as eachhigh-efficiency heat exchanger has a close temperature of approach witha counter-current design to minimize the amount of cooling waternecessary and to maximize the amount of steam recovery. Furthercondensation is achieved through the introduction of an additionalcooling stream to the high-efficiency heat exchanger so as to effectthermal contact with the product bearing stream and increase the amountof product condensed from the product bearing stream. Further, contraryto other references which state that the interstage removal of methanolby condensation is not practical or economical, such as K R. Westerterp,New Methanol Processes, “Energy Efficiency in Process Technology” Ed. P.A. Pilarvachi, Elsevier Applied Science, 1993, pp. 1142-1153, 1147, theuse of the high-efficiency plate type or extended surface heatexchangers to provide interstage methanol condensation operates in botha practical and economically viable manner within the apparatus andmethod of this invention.

According to another aspect of the present invention, the variousreactor sets of the cascade system are formed as reaction zonesintegrated within a single reactor vessel, such that only the reactorvessel and appropriate inlet and outlet fittings on the vessel need tobe constructed to withstand the temperatures and pressures necessary forthe methanol production reaction. The reaction products from eachreaction zone are passed through the high-efficiency heat exchangerswhich are also formed in a block-like configuration positioned andconnected between each reaction zone in the reactor vessel, and aseparator zone located in a separate separator vessel constructedsimilarly to the reactor vessel. The construction of the variousreaction zones within the reactor vessel and the various separator zoneswithin the separation vessel greatly reduces the cost of the materialsnecessary to construct the various vessels, as the pressuredifferentials between the respective zones in each of the reactor vesseland the separation vessel are minimal. This eliminates the need forconstructing individual walls between the various zones of materialscapable of withstanding the high pressure differentials between reactionpressure and atmospheric pressure that would otherwise be encountered.

According to still a further aspect of the present invention, animproved multiple or cascade reactor set type methanol production systemis provided in which methanol is initially produced through the use ofany suitable methanol production system, such as a conventional methanolsynthesis loop with a recycle compressor. The purge gas stream from themethanol production system is subsequently directed through a cascadesystem of three or more reactor sets formed according to the presentinvention that further react the unreacted components of the purge gasstream to form additional methanol. The reactors are constructedseparately, or as part of a single vessel with a separating walldesigned to contain the differential pressure between the reactors andprovide access between reactors to aid filling of the individual bedswith catalyst. In each reactor set a single multi-stream heat exchangeris used to recover high grade heat, to effect feed/effluent heatexchange, and to achieve condensation of methanol by further coolingwith a cooling medium such as water. In addition, the compactmulti-stream heat exchangers are arranged in such a manner alongside thereactor set so as to minimize the amount of connections between thecascade reactor set and the synthesis loop.

According to still another aspect of the present invention, the multiplereactor set utilizing the high-efficiency plate-type or extended surfaceheat exchangers can be utilized as a stand-alone stationary or mobilesystem and/or as an add-on to an existing recycle loop reactor set or toan existing cascade reactor set to further increase the percentconversion of methanol from these pre-existing reactors, or maintain theoverall conversion of the modified process while relaxing theeffectiveness of the recycle process through, for example, a reducedrecycle rate.

Numerous additional aspects, features, and advantages of the presentinvention will be made apparent from the following detailed descriptiontogether with the drawing figures.

BRIEF DESCRIPTION OF THE DRAWINGS

The drawings illustrate the best mode currently contemplated ofpracticing the present invention.

In the drawings:

FIG. 1 is a schematic view of a prior art recycle loop reactor system;

FIG. 2 is a schematic view of a prior art multiple reactor set orcascade system;

FIG. 3 is a schematic view of the multiple reactor set of FIG. 2including a number of interstage feed/effluent heat exchangersconstructed according to the present invention;

FIG. 4 is schematic view of a second embodiment of the reactor set ofFIG. 3, in which the separate reactors, heat exchangers and separatorsare formed as zones disposed within a single vessel;

FIG. 5 are isometric views of header constructions for the feed/effluentheat exchangers of the reactor of FIG. 4;

FIG. 6 is a schematic view of a weir discharge system for the separatorsof the reactor of FIG. 4; and

FIG. 7 is a schematic view of the reactor set of FIG. 3 attached to thepurge gas stream of a methanol production reactor system.

DETAILED DESCRIPTION OF THE PREFERRED EMBODIMENTS

With reference now to the drawing figures in which like referencenumerals designate like parts throughout the disclosure, a methanolproduction system that can be utilized with or replaced by the presentinvention is indicated generally at 100 in FIG. 1. The system 100 shownis a conventional methanol synthesis loop that receives a stream 160 ofsyn-gas from a reformer 120, converts a portion of the syn-gas tomethanol in a reactor 140, and discharges a combined stream 180 ofunreacted syn-gas and methanol to a condenser or separator 200. Thecombined stream 180 is separated in the separator 200 into a methanolstream 220 that is collected, and a recycle stream 240 that is directedfrom the separator 200 to a recycle compressor 260. The compressor 260compresses the gas in the recycle stream 240 and directs therecompressed gas stream 280 back to the reactor 140. However, a portionof the recycle stream 240 entering the compressor 280 is diverted as apurge gas stream 300 prior to recompression.

In addition, another prior art methanol production system that iscapable of use with the present invention is the multiple reactor set orcascade system 320 shown in FIG. 2. Similar to the recycle system 100,the cascade system 320 also includes the reformer 120 that forms anddirects a stream 160 of syn-gas to a first reactor 140 that converts aportion of the stream 160 to methanol, forming a combined stream 180that exits the reactor 140. This combined stream 180 subsequently entersa first separator 200 where it is separated into a methanol stream 220and a first stream 340 of excess syn-gas. This excess syn-gas stream 340is then directed to a second reactor 360 for further reaction of thestream into methanol. The resulting second combined stream 380 is thenpassed to a second separator 400, where a second methanol stream 220′ isformed and combined with stream 220, and a second excess syn-gas stream340′ is passed to a third reactor 360′. The third reactor 360′ formsfrom the second excess stream 340′ a third combined stream 380′ thatpasses into a third separator 400′ and which exits as a third methanolstream 220″ and a third excess syn-gas stream 340″. The third methanolstream 220″ is combined with streams 220 and 220′, and the third excessstream 340″ is passed to a fourth reactor 360″. The reactor 360″transforms the excess stream 340″ into a fourth combined stream 380″that enters a fourth separator 400″. The fourth separator 400″ separatesthe combined stream 380″ into a fourth methanol stream 220″′ and a purgegas stream 420.

A multiple reactor set methanol production system constructed accordingto the present invention is indicated generally at 10 in FIG. 3. Thesystem 10 receives a stream 16 of syn-gas from a reformer (not shown),converts the syn-gas to methanol in stages, and ultimately discharges astream 21 of purge gas and a combined methanol stream 25. Significantly,it lacks a recycle compressor yet still operates efficiently, permittingit to be used economically in relatively small-scale applicationsproducing less than 2,500 tons of methanol per day, and moreparticularly on the order of less than 1,500 or less than 1,000 tons ofmethanol per day. It could, however, be scaled up for larger scaleoperation as well or reduced in size without significant detriment.

The syn-gas typically contains approximately 66 mol % hydrogen, 20 mol %carbon monoxide, 9 mol % carbon dioxide, and 2 mol % methane. It wouldalso contain any nitrogen that was present in the methane originally fedto the reformer where the syn-gas is made. The actual composition willdepend on the pressure and temperature used in the reforming, the methodof reforming (steam reforming, autothermal etc.) and whether there wasany carbon dioxide added to, or present in the methane stream fed to thereformer.

The system 10 includes a number of reactor sets 11, 13, 15, 17 locatedin series such that each downstream reactor set receives the effluentfrom the immediately upstream reactor set as a feed stream, furtherconveys the feed stream, and discharges a condensed methanol stream 24and an effluent stream 20. The methanol streams 24 are combined to formcombined stream 25. The effluent stream from the downstream-most reactorset forms the purge gas stream 21.

Still referring to FIG. 3, each of the reactor sets 11, 13, 15, 17includes a reactor 12, a separator 18, and a feed/effluent heatexchanger 22. Each of the reactors 12 receives a feed stream feed 16 or20 and discharges a product stream 14. The product stream 14 from eachof the reactors 12 is directed through a corresponding heat exchanger22, where it is cooled by heat exchange from the feed stream 16 or 20for the reactor of that set to condense the methanol from the productstream. The stream 14 is additionally cooled by the introduction ofcooling utility within the structure of the feed-effluent exchanger 22such that thermal contact is achieved. The product stream 14 of eachreactor set 11, 13, 15, 17 is then directed to a corresponding separator18 of any suitable type which, in turn, provides a return stream 20 tothe reactor 12 of the next downstream reactor set. As mentioned above,the return stream 20 from the separator 18 of the final reactor set 15is discharged as a purge gas stream 21.

The reactors 12 can be selected to be any suitable type of reactor foruse in a methanol production reaction as are known in the art.Preferably each reactor consists of multiple adiabatic beds with coolingof the process fluid after each bed such that steam is produced to beused as a utility. Only two adiabatic beds per reactor are sufficient toenable the correct operation of the process. Suitable alternate reactorscould be conceived by comparison with conventional methanol reactorssuch as steam jacketed tubes (or Lurgi reactor), a tube cooled reactoravailable from Davy-Synetix, an adiabatic bed reactor with more than twobeds, or perhaps spherical or radial geometry multiple adiabatic bedreactors, among others. Each reactor 12 contains a methanol synthesiscatalyst such as one comprising a reduced zinc oxide/copper oxidemixture. The conversion typically takes place at 40-100 bars andtemperatures in excess of 200° C., typically at 220-280° C. but notexceeding 310° C. Typically the product stream 14 will containapproximately 5 mol % methanol. Higher conversion rates can be achievedat lower temperatures but at the expense of larger catalyst volumes.

The heat exchangers 22 are preferably selected to be plate-type heatexchangers, such as a diffusion bonded printed circuit heat exchangermanufactured by Heatric of Dorset, UK, or extended surface heatexchangers, such as a diffusion bonded plate fin heat exchangeravailable from Heatric, or a brazed plate fin heat exchanger availablefrom Chart Industries of Bracknell, UK, a spiral wound heat exchanger,or other suitable stacked plate heat exchangers, as opposed to prior artshell and tube heat exchangers. These types of heat exchangers arepreferred because plate-type or extended surface heat exchangers 22 arecapable of providing a close temperature of approach to the fluidstreams passing through the exchangers 22, such that the product streamsexit the exchangers 22 within five (5) degrees C. of one another. Thiseffectively minimizes the amount of cooling water required and maximizesthe steam recovery from the heat exchangers 22, such that these types ofheat exchangers have an effectiveness greater than 7 NTU.

The heat exchangers 22 also are capable of multi-stream and/orcountercurrent operation such that multiple heat transfer operations cantake place within a single heat exchanger 22. Specifically, the heatexchangers 22 effectively cool the methanol component within each of theproduct streams 14 to condense the methanol in the product streams 14and enable removal of the methanol within the separators 18 in a highlyeconomic manner. An additional cold utility stream is introduced intothe heat exchanger to maximize methanol condensation. In addition, theheat retained by the product streams 14 is effectively utilized toelevate the temperature of the feed stream 16 and the return streams 20prior to these streams 16 and 20 entering the reactors 12. Methanoltypically begins to condense at 110° C., depending on concentration andreaction pressure. For efficient removal of methanol (>75%) where theexit concentration is typical (5%), then the heat exchangers 22 arepreferably configured to cool the product streams 16 to below 60° C. bythe incorporation of a cooling utility stream into the feed/effluentexchange reactor. Most preferably, the reactor return stream 20 can passthrough a single plate style heat exchanger 22 where, in a firstsection, the gas is cooled with a high temperature coolant such aspressurized water at 200-250 deg C. In a second section of the heatexchanger 22, the reactor gases are then cooled by thermal contact withthe reactor inlet stream 16. Finally, in a third section, the reactorgases are cooled with cooling water to condense the water.

While the system 10 shown in FIG. 3 illustrates four reactors sets 11,13, 15, 17, the number of reactor sets and constituent components ofeach reactor set can be varied as desired. For instance, optimaloperation of the illustrated embodiment with four reactors 12 requires atight composition control of the syn-gas in the feed stream 16 in orderto keep the stoichiometric ratio [(H2−CO₂)/(CO+CO₂)] between 2:1 and3:1, and preferably between 2.1-2.2/1 in order to achieve the 95 percentconversion of the syn-gas to methanol required for an economicallyviable process. However, by adding additional reactor sets, up to ten(10), a system would capable of producing acceptable overall syn-gasconversions (i.e., in excess of 95 percent for CO_(BS)), or in excess of90 percent for H₂) with a wider range of feed gas compositions havingratios between 3:1 and 2:1, and/or for feed gases with changingcompositions, such as are present at CO₂ stimulated gas fields.

Referring now to FIG. 4, in a more specific embodiment of the invention,the system 10′ includes a reactor vessel 27 that defines a number ofreaction zones 26 therein, each of which contains the reactor 12 for acorresponding reactor set 11′, 13′, 15′, etc. The system 10′ alsoincludes a separator vessel 30 which defines a number of separationzones 32 therein for each of the reactor sets 11′, 13′, 15′, etc. Thevessels 27 and 30 are constructed in a manner which allows the vessels27 and 30 to withstand the elevated temperature (200° C. to 310° C.) andpressure (40-100 bars) required for the methanol production reaction. Toseparate the various reaction zones 26, dividing walls 28 are disposedbetween adjacent zones 26. Because all of the reaction zones 26 aredisposed within the vessel 27, the only pressure differential betweenthe zones 26 is the pressure drop between the process streams flowinginto and out of successive reaction zones 26, which is typically around0.2-2.0 bars. Thus, the dividing walls 28 are constructed of materialsthat only need to be able to withstand a pressure differential of around0.2-2.0 bars, which is much less costly than the materials forming thevessel 27, and the walls 28 can be of a simple welded construction. Easeof construction may result in a design that uses more than one reactorvessel to accommodate the multiple reactor zone and more than oneseparation vessel to accommodate the methanol separators. Additionally,access to each of the zones 26 can be provided through the walls 28 atan internal access point (not shown) capable of withstanding 2 bars ofpressure, instead of an external access point (not shown) requiring anforty (40) to one hundred (100) bar design pressure. Also, with theinclusion of the reactors 12, heat exchangers 22 and separators 18 inthe various vessels 27 and 30, and heat exchanger block 40, to bediscussed, the system 10 can be operated at the same pressure as thereformer (not shown) that supplies the syn-gas to the system 10, thuseliminating the need for syn-gas compression prior to being fed to thesystem 10.

Similarly, the specific embodiment of the system 10′ also includes aseparation vessel 30 that defines a number of separation zones 32 orknock out pots, each of which contains the heat exchanger 22 andseparator 18 for a corresponding reactor set 11′, 13′, 15′. Theseparation zones 32 are separated from one another by dividing walls 34.Again, because the pressure drop between adjacent separation zones 32 isvery low, e.g., less than two (2) bars, the dividing walls 34 can beconstructed of materials similar to walls 28 and much less costly thanthe materials utilized for the construction of the separation vessel 30.Additionally, as shown in FIG. 6, due to the small pressure drop betweenzones 32, a weir system 36 can be utilized that connects each of thezones 32 and allows the condensed methanol stream 24 to flow from aseparation zone 32 downwardly along a liquid drain 38 into an adjacentseparation zone 32 under the influence of the pressure differentialsbetween the separation zones 32. The liquid drain 38 and weir system 36thereby allow the methanol stream 24 to flow between the respectiveseparation zones 32 or knock-out pots to form the combined stream 25without the need for an active level control of the methanol levelwithin the separation zones 32 and still maintaining a gas seal. In itssimplest construction, the zones 32 can be formed with a mixed phaseinlet (the product stream 14), an upper gas outlet (the return stream20), and a lower fluid outlet (the methanol stream 24) when the liquidcan directly settle out of the gas in the mixed phase by gravity.

Looking at FIGS. 4 and 5, the heat exchangers 22 can also beincorporated into a single block 40 where each of the exchangers 22 arepositioned to align the entry and exit points (not shown) for thecooling fluid flowing through the exchangers 22 such that they can beconnected to a header 42 that, in turn, is operably connected to piping44 to distribute a cooling fluid flow into the inlet 45 for each of therespective heat exchangers 22. As the heat exchangers 22 are formed ofextended surface or plate-type heat exchangers, such as printed circuitheat exchangers, a single header 42 can be utilized for a single fluidsupplied to each of the heat exchangers 22 in order to provide thecooling water and/or heated steam to the exchangers 22. A similar typeof header (not shown) but with dividing walls (not shown) can also bedisposed on each exchanger 22 in the block 40 that is configured tofunction as a reaction zone that replaces the zones 26 in the vessel 27by placing a reaction catalyst in the header, which receives the feedstream 14 or one of the return streams 20 from the separation zones 32.The header allows the methanol conversion reaction to take place withinthe header and subsequently redirects the product stream 16 with themethanol and excess syn-gas back into the heat exchanger 22 to which theheader is connected.

In the particular embodiments in FIGS. 4-6 where the methanol productionsystem 10′ includes a reactor vessel 27 with multiple reaction zones 26,a block 40 with heat exchangers 22 and a separation vessel 30 withseparation zones 32, the system 10′ can be configured to be constructedeither as a mobile unit, or a fixed unit that has the capacity tosynthesize from 1 to 500 metric tons per day of methanol. Also, inaddition to the use of the system 10′ as a stand-alone unit, the system10′ can be connected to the purge gas stream of a recycle loop reactoror any other methanol production system to increase the conversion ofthe recycle loop reactor or multiple reactor set by using the purge gasas the feed stream 16 for the reactor set 10′. This use is especiallyadvantageous where the reactor set system 10′ is formed with the reactorvessel 27, heat exchanger block 40, and the separation vessel 30.

Looking now at FIG. 7, another embodiment of the present invention isillustrated in which the system 10 is utilized as a methanol recoverysystem 52 that is operably connected to the purge gas stream 50 of arecycle compressor 260 of a recycle loop system 100, similarly to thatdescribed previously for the system 10′. The system 52 includes a numberof reactors 54, 54′, 54″ operably connected to one another, and whichpreferably are formed as simple adiabatic reactors. The purge gas stream50 exiting the recycle compressor 260 initially passes through the afirst heat exchanger 56 to preheat the purge gas stream 50 prior toentering the first reactor 54. The purge gas stream 50 is heated by afirst reactor product stream 58 that exits the first reactor 54 andpasses through the first heat exchanger 56 to thermally contact andraise the temperature of the purge gas stream 50. Simultaneously, thepurge gas stream 50 lowers the temperature of the first product stream58 which consists of methanol and still unreacted purge gas. Thisnow-cooled first product stream 58 then passes from the first heatexchanger 56 to a first separator 60 whereby the product stream 58 isfurther cooled to produce a first methanol stream 62 and a firstunreacted purge gas stream 64. The first methanol stream 62 is collectedfrom the first separator 60 to form a methanol product stream 90, whilethe first unreacted purge gas stream 64 is directed to a second heatexchanger 56′ in order to cool and be heated by a second product stream58′ coming from the second reactor 54′ of the methanol recovery system52 in the same manner as described regarding first heat exchanger 56.Additionally, the second product stream 58′ exiting the second reactor54′ is processed by the second heat exchanger 56′ and a second separator60′ in a manner similar to the product stream 58 exiting the firstreactor 54 in order to generate a second methanol stream 62′ that iscollected from the second separator 60′ and added to the methanolproduct stream 90, and a second unreacted purge gas stream 64′.

The second unreacted purge gas stream 64′ is directed from the secondseparator 60′ to a third reactor 54″ through a third heat exchanger 56″in the same manner as described previously regarding the passage ofpurge gas stream 50 and first unreacted purge gas stream 54 through heatexchangers 56 and 56′. The third reactor 54″ uses the second unreactedpurge gas stream 64″ to generate a third product stream 58″ that isdirected through the heat exchanger 56″ to a third separator 60″ thatgenerates a collectible third methanol stream 62″ that is added tomethanol production stream 90, and a third purge gas stream 54″ which isdischarged from the system 52.

The number of reactors 54, 54′ and 54″ can be varied as necessary fromone to any number required for the desired methanol conversion, and canbe selected to be any suitable type of reactor for use in a methanolproduction reaction as are known in the art. Preferably each reactorconsists of a simple adiabatic reactor, and most preferably withmultiple adiabatic beds, with cooling of the process fluid after eachbed such that steam is produced to be used as a utility. Only twoadiabatic beds per reactor are sufficient to enable the correctoperation of the process. Suitable alternate reactors could be selectedto be similar to those described previously as alternatives for thereactors 12 utilized in the system 10.

The heat exchangers 56, 56′ and 56″ are constructed similarly to theheat exchangers 22 discussed previously, and are preferably selected tobe plate-type heat exchangers, such as a diffusion bonded printedcircuit heat exchanger, or extended surface heat exchangers, such as adiffusion bonded plate fin heat exchanger or a brazed plate fin heatexchanger, as opposed to prior art shell and tube heat exchangers. Thesetypes of heat exchangers are preferred for the same reasons describedwith regard to the heat exchangers 22 utilized in the system 10, namelydue to their ability to provide a close temperature of approach to thefluid streams passing through the exchangers 56, 56′ and 56″, and theability of the exchangers to function in a multi-stream and/or countercurrent manner. This effectively minimizes the amount of cooling waterrequired and maximizes the steam recovery from the heat exchangers 56,56′ and 56″. As a result, the heat exchangers 56, 56′ and 56″ alsoeffectively cool the methanol component within each of the productstreams 58, 58′ and 58″ to condense the methanol in the product streams58, 58′ and 58″ and enable removal of the methanol within the separators60, 60′ and 60″ in a highly economic manner.

Further, the reactor product streams 58, 58′ and 58″ can be cooled by autility stream (not shown) prior to entering the heat exchangers 56, 56′and 56″. Significantly, it lacks a recycle compressor yet still operatesefficiently, permitting it to be used economically in relativelysmall-scale applications producing on the order of 1,000 to 1,500 tonsof methanol per day. It could, however, be scaled up for larger scaleoperation as well or reduced in size without significant detriment.

Additionally, the methanol recovery system 52 can include combinationelements (not shown) which function as both each of the heat exchangers56, 56′ and 56″ and separators 60, 60′ and 60″ to minimize the number ofcomponents utilized in the methanol recovery system 52.

Further, the methanol recovery system 52 can be utilized as astand-alone stationary or mobile system and/or as an add-on to anexisting recycle loop reactor system 100 or to an existing cascadereactor system 320 to further increase the percent conversion ofmethanol from these pre-existing reactors, or maintain the overallconversion of the modified process while relaxing the effectiveness ofthe recycle process through for example reduced recycle rate. Also, thesystem 52 can be used with systems 10 that produce other compoundsformed via equilibrium limited reactions, such as higher alcohols ordimethyl ether, among others.

Various alternatives are contemplated as being within the scope of thefollowing claims particularly pointing and distinctly claiming thesubject matter regard as the invention.

1. A method for the synthesis of a chemical compound formed by anequilibrium limited reaction, the method comprising the steps of: a)providing a reactor set including a first heat exchanger, and reactoroperably connected to the first heat exchanger and a second heatexchanger operably connected to the reactor; b) directing a feed streamthrough the first heat exchanger formed as a plate-type or extendedsurface heat exchanger to preheat the feed stream; c) reacting the feedstream within the reactor; d) discharging a reaction product stream fromthe reactor through the first heat exchanger to preheat the incomingfeed stream and to cool the reaction product stream; e) thermallycontacting the reaction product stream with a cooling utility stream toform a condensed product within the reaction product stream; f)separating the condensed product from the reaction product; g) directingthe reaction product stream through the second heat exchanger formed asa plate-type or extended surface heat exchanger to preheat the reactionproduct stream; and h) repeating steps b-g to achieve the desiredproduct conversion.
 2. The method of claim 1 where the feed stream iscomprised of hydrogen and carbon oxide gases.
 3. The method of claim 2wherein the feed stream is a purge gas stream from a recycle stream of achemical production reactor utilizing a steam reformer, and wherein thestep of directing the feed gas stream comprises directing the purge gasstream through the first heat exchanger.
 4. The method of claim 1wherein the first heat exchanger is formed with a first portion forthermally contacting the feed stream with the reaction product stream,and a second portion downstream from the first portion for thermallycontacting the reaction product stream with the cooling utility, andwherein the step of thermally contacting the reaction product streamwith the cooling utility takes place in the first heat exchanger.
 5. Themethod of claim 4 wherein the first heat exchanger includes a thirdportion downstream from the second portion for separating the condensedproduct from the reaction product stream, and wherein the step ofseparating the condensed product from the reaction product stream takesplace in the third portion of the first heat exchanger.
 6. The method ofclaim 1 further comprising a separator operably connected between thefirst heat exchanger and the second heat exchanger, and wherein the stepof separating the condensed product from the reaction product streamcomprises passing the condensed product and the reaction product streamthrough the separator.
 7. The method of claim 1 wherein the chemicalcompound is selected from the group consisting of: methanol, dimethylether and mixtures thereof.
 8. A product synthesis apparatus for theproduction of a product formed by an equilibrium-limited reaction, theapparatus comprising: a) a reactor vessel adapted to withstand theoperating temperatures and pressures of the product-forming reaction,the reactor vessel including a number of reaction zones formed withinthe reactor vessel by dividing walls extending across the reactor vesselto define the reaction zones, the dividing walls configured to withstandtemperature and pressure differentials across each adjacent reactionzones; b) a separation vessel adapted to withstand the operatingtemperatures and pressures of the product-forming reaction; and c) aheat exchanger operably connected to at least one of the reaction zonesand to the separation vessel.
 9. The apparatus of claim 8 wherein thedividing walls are configured to withstand pressure differentials ofbetween 0.2 to 2.0 bars across each adjacent reaction zones.
 10. Aproduct synthesis apparatus for the production of a product formed by anequilibrium-limited reaction, the apparatus comprising: a) a reactorvessel adapted to withstand the operating temperatures and pressures ofthe product-forming reaction; b) a separation vessel adapted towithstand the operating temperatures and pressures of theproduct-forming reaction, the separation vessel including a number ofseparation zones formed within the separation vessel by dividing wallsextending across the separation vessel to define the separation zones,each dividing wall configured to withstand the temperature and pressuredifferentials across adjacent separation zones; and c) a heat exchangeroperably connected to the reactor vessel and to at least two of theseparation zones.
 11. A method for condensing methanol out of a reactionproduct stream from a reactor of a methanol production system, themethod comprising thermally contacting the reaction product stream witha feed stream for the reactor in a plate type or extended surface heatexchanger to condense the methanol in the reaction product stream.
 12. Amethod for producing methanol comprising: a) providing a multiplereactor set including a number of reactors for converting a feed gasstream formed of a syn-gas created in a steam or autothermal reformerinto a reaction product stream, a number of feed/effluent heatexchangers connected to the reactors and configured to preheat the feedgas stream, to cool the reaction product stream, and to condense thereaction product stream into a methanol stream and an excess feed gasstream, and a number of separators connected to the heat exchangersopposite the reactors for separating the methanol stream from the excessfeed gas stream; and b) condensing methanol from the reaction productstream at a rate of less than 2,500 tpd.
 13. The method of claim 12further comprising condensing methanol from the reaction product streamat a rate of less than 1,500 tpd.
 14. The method of claim 12 furthercomprising condensing methanol from the reaction product stream at arate of less than 1,000 tpd.
 15. A method for producing a compound froma waste gas stream of a production system for the compound that employsa steam or autothermal reformer comprising the step of passing the wastestream through a multiple reactor set including a number of reactors forconverting a feed gas stream formed of the waste gas stream into areaction product stream, a number of feed/effluent heat exchangersconnected to the reactors and configured to preheat the feed gas stream,to cool the reaction product stream, and to condense the reactionproduct stream into a product compound stream and an excess feed gasstream, and a number of separators connected to the heat exchangersopposite the reactors for separating the product compound stream fromthe excess feed gas stream, wherein the multiple reactor set does notinclude a compressor.
 16. The method of claim 15 wherein the wherein thefeed stream is a purge gas stream from a chemical production reactorutilizing an autothermal reformer.